Engineering
Master Thesis
Kinetic modelling of fluidized bed biomass gasification
integrated with steam electrolysis for synthetic
methane production
Supervisors
Prof. Massimo Santarelli Ing. Emanuele Giglio
Candidate
Antonio De Padova
A process configuration for the production of synthetic natural gas (SNG) starting from woody biomass is proposed and analyzed.
A kinetic rate model was developed to simulate the gasification of biomass in a bubbling fluidized bed (BFB) gasifier with a mixture of oxygen and steam as gasifying agent. The model has been then validated comparing its pre- dictions with the experimental results reported in the literature and with the predictions of another model, used as a reference.
The gasification stage performance, in terms of outlet stream composition and temperature, have been then analyzed by varying some parameters like equivalence ratio (ER) and steam-to-biomass ratio (SBR). Proper values of ER and SBR have been set to maximize hydrogen content within the produced syngas.
A solid oxide electrolysis cell (SOEC) system for water splitting has been then modeled. Hydrogen produce at cathode side is mixed with the syngas exiting the gasification unit, to reach the stoichiometric syngas composition for the subsequent methanation reaction, where carbon monoxide and car- bon dioxide are hydrogenated in a catalytic reactor to produce synthetic methane. Oxygen-steam mixture from anode outlet is exploited as gasifying agent within biomass gasification unit.
A simple approach for the syngas cleaning was used to take into account all the downstream equipment used to reduce the contaminants (e.g., tars, ammonia, hydrogen sulfide and hydrogen chloride) content in the raw syngas before the methanation unit.
The whole process has been modeled and analyzed using ASPEN Plus pro- cess simulator integrated with an external Fortran subroutine to represent properly the gasification stage in terms of hydrodynamics and reaction kinet- ics.
I take this opportunity to express my gratitude to my supervisors, Prof. Mas- simo Santarelli, for the chance of developing this thesis, and Ing. Emanuele Giglio, for his guidance, patience and availability in any moment during these months.
I wish to express my gratitude to my parents for their constant encourage- ment and unconditional love.
Thanks to my family and my friends for their support and their presence both in the hard moments and in the joyful ones.
Without all of you I would not be who I am today.
"Happiness is only real when shared."
List of Figures 6
List of Tables 8
List of Symbols 9
1 Introduction 11
1.1 Research objective . . . 12
2 Technology review 15 2.1 Biomass . . . 15
2.1.1 Biomass composition . . . 16
2.1.2 Biomass conversion . . . 17
2.2 External drying . . . 17
2.3 Gasification . . . 18
2.3.1 Types of gasifiers . . . 21
2.3.2 Equivalence Ratio (ER) and Steam-to-Biomass Ratio (SBR) 26 2.3.3 Tars formation . . . 27
2.4 Electrolysis . . . 29
2.4.1 Polarization curve . . . 32
2.4.2 Cell efficiency . . . 34
2.4.3 Solid Oxide Electrolytic Cell (SOEC) . . . 35
2.4.4 SOEC materials . . . 36
2.5 Syngas clean-up system . . . 37
2.5.1 Cold gas clean-up . . . 37
2.5.2 Hot gas clean-up . . . 38
2.6 Methanation . . . 39
3 Process modelling 41 3.1 General aspects . . . 41
3.2 External drying section . . . 44
3.3 Gasification section . . . 46
3.3.3 Gasifier modelling . . . 52
3.4 Solid Oxide Electrolytic Cell (SOEC) . . . 62
3.5 Syngas cleaning unit . . . 63
3.6 Methanation unit . . . 65
4 Simulation results and process integration 67 4.1 Sensitivity analysis . . . 67
4.1.1 Effect of ER on syngas composition . . . 68
4.1.2 Effect of SBR on syngas composition . . . 69
4.1.3 Effect of temperature on syngas composition . . . 70
4.2 Parameters dependency . . . 71
4.2.1 Effect of ER on Tgas . . . 72
4.2.2 Effect of ER on syngas molar composition . . . 73
4.2.3 Effect of SBR on Tgas . . . 74
4.2.4 Effect of SBR on syngas molar composition . . . 75
4.3 System optimization . . . 76
4.4 Tar production estimation . . . 82
4.4.1 Effect of ER on tar yield and concentration . . . 82
4.4.2 Effect of SBR on tar yield and concentration . . . 85
4.5 Electrolysis integration . . . 88
4.6 Methanation performances . . . 89
4.7 Possible thermal integration . . . 89
5 Conclusions 91 5.1 Recommendations for future works . . . 93
Bibliography 94
1.1 Total Energy Supply (TES) by source in 1971 and 2018 [1] . . . . 11
2.1 Devolatilization products . . . 19
2.2 Fixed bed gasifiers: 1. updraft gasifier, 2. downdraft gasifier, 3. crossdraft gasifier [16]. . . 22
2.3 Fluidized bed gasifiers: 1. Bubbling fluidized bed gasifier, 2. Circulating fluidized bed gasifier [22]. . . 24
2.4 Entrained bed gasifier [23] . . . 25
2.5 Distribution of tar species as a temperature function [21] . . . 28
2.6 Conceptual design of a SOEC [12] . . . 29
2.7 Total, thermal and electrical energy demand for an ideal electrol- ysis process as a function of T [27] . . . 31
2.8 Overview of characteristic I-U curves of SOECs from literature [27] . . . 33
2.9 Simplified layout of a SOE system [27] . . . 36
3.1 Simplified process flowchart . . . 43
3.2 External drying unit flowsheet . . . 44
3.3 Gasification unit flowsheet . . . 46
3.4 SOEC unit flowsheet . . . 62
3.5 Cleaning unit flowsheet . . . 64
3.6 Methanation unit flowsheet . . . 65
4.1 Dry syngas molar composition as a function of ER at T = 850°C and SBR = 0.35 . . . 68
4.2 Dry syngas molar composition as a function of SBR at T = 850°C and ER = 0.27 . . . 69
4.3 Dry syngas molar composition as a function of Tgas with ER = 0.27 and SBR = 0.35 . . . 70
4.4 Gasifier control volume . . . 71
4.5 Tgas as a function of ER . . . 72
gas
4.8 Molar composition of the syngas for different values of ER and Tgas75 4.9 Gasification temperature as a function of ER for different values
of SBR . . . 78 4.10FEED parameter as a function of ER for different values of SBR . 78 4.11Gasification temperature as a function of SBR for different values
of ER . . . 81 4.12FEED parameter as a function of ER for different values of SBR . 81 4.13Tar concentration as a function of ER for different values of SBR. 84 4.14Tar yield as a function of ER for different values of SBR. . . 84 4.15Tar concentration as a function of SBR for different values of ER. 87 4.16Tar yield as a function of SBR for different values of ER. . . 87 4.17Control volume of the electrolysis unit in the optimal configuration 88 5.1 Simplified flowchart with main model results . . . 92
3.1 Gasification chemical reaction rates . . . 58 3.2 Biomass composition used for model validation . . . 59 3.3 Comparison of model results with Dang et al. [11] results and
Campoy et al. [11] experimental results, expressed as composi- tion of the non-condensable gaseous phase on a dry N2-free basis 60 3.4 Pine#126 biomass composition [46] . . . 61 3.5 Geometrical parameters for the gasifier with steam and air (Case
1) and with steam and oxygen (Case 2) as gasifying agent . . . 61 4.1 Evaluation of the molar composition of the syngas and FEED pa-
rameter as a function of ER for different values of SBR . . . 77 4.2 Evaluation of the molar composition of the syngas and FEED pa-
rameter as a function of SBR for different values of ER. . . 80 4.3 Tar yield and concentration as a function of ER for different val-
ues of SBR . . . 83 4.4 Tar molar composition as a function of SBR for different values
of ER . . . 83 4.5 Tar mass flow and yield as a function of SBR for different values
of ER . . . 86 4.6 Tar molar composition as a function of SBR for different values
of ER . . . 86 4.7 Methanation unit streams molar compositions . . . 89 4.8 Heat produced or required in the optimal system configuration . 90
Symbol Definition
∆h0 Standard enthalpy of formation [kJ/mol]
˙n Molar flow rate [kmol/s]
∆G Change in Gibbs free energy [J]
∆H Overall energy demand of reaction [J]
∆Q Heat of reaction [J]
T Temperature [°C] or [K]
∆S Entropy variation [J/K]
U Potential [V]
z Number of electrones
F Faraday constant
η Overvoltage [V]
j Current density [A/m2]
S Active surface area [m2]
nC Number of cells
ηF Faraday efficiency
V˙ Volumetric flow rate [Nm3/h]
Pel Electrical power [W]
ES Specific energy consumption [kWh]
˙
m Mass flow rate [kg/s]
d Diameter [m]
ρ Density [kg/m3]
Ar Archimedes number
g Gravitational acceleration [m/s2]
µ Gas viscosity [Pa·s]
umf Minimum fluidization velocity [m/s]
εmf Voidage in the emulsion phase [-]
db0 Initial bubble diameter [m]
N Number of orifices of the distributor plate
b
δb Fraction of bed occupied by bubbles [-]
vbedave Average bed voidage [-]
εf Volume occupied by the emulsion phase hinc Incremental height [m]
C Molar concentration [kmol/m3]
p Pressure [Pa]
Rgas Ideal gas constant [kJ/(kmol·K)]
X Carbon conversion factor
yi Molar fraction of the i-th species h Specific mass enthalpy [kJ/kg]
Abbreviation Definition
SOEC Solid Oxide Electrolytic Cell
SNG Synthetic Natural Gas
BFB Bubbling Fluidized Bed
ER Equivalence Ratio
SBR Steam-to-Biomass Ratio
DME Dimethyl ether
NG Natural Gas
RMSE Round Mean Square Error
UNFCCC United Nations Framework Conventions on Climate Change
MOIST Moisture
ULT Ultimate Analysis
PROX Proximate Analysis
RR Reactant Ratio
RU Reactant Utilization
OCV Open Circuit Voltage
ASR Area Specific Resistance
HHV Higher Heating Value
LHV Lower Heating Value
YSZ Yttrium Stabilized Zirconia
LSM Lanthanum Strontium Manganite
FEED Feed Gas Module
FACT Correction factor used to refer to dry biomass mass flow
WGS Water Gas Shift
RWGS Reverse Water Gas Shift
SMR Steam Methane Reforming
Introduction
From 1971 to 2018, the World Total Energy Supply (TES) grown very fast:
from 5,519 Mtoe to 14,282 Mtoe, it became almost three times bigger (Fig- ure 1.1) [1]. However, in this scenario, conventional fossil fuels still play a dominant role in the energy mix. Due to a decreasing availability of the fossil fuels and to a rising attention to themes like climate change, the re- newable energy sources are starting to play a key-role. Their contribution to the energy mix is around the 11%, considering the sum of biofuels and other renewable energy sources.
Figure 1.1: Total Energy Supply (TES) by source in 1971 and 2018 [1]
Biomass is for sure one of the most important renewable energy sources,
neutral resource: burning fossil fuels converts “old” biomass into “new” CO2, which is injected in the environment contributing to the “greenhouse effect”, while biomass combustion produces the same CO2 that it absorbed during its life cycle. Furthermore, biomass does not take millions of years to grow and to develop again, as it happens for carbon or oil, and this is the reason why it is considered a renewable resource [2, 3].
Biomass is exploited with different thermo-chemical processes, mainly combustion, pyrolysis and gasification. With gasification, biomass molecules are broken down to produce gaseous products, named syngas. Syngas can be directly used with energetic purposes or used for the production of secondary products of interest, named biofuels, like dimethyl ether (DME), Fischer- Tropsch products, methanol and synthetic natural gas (SNG). Generally bio- fuels can be used with a series of advantages with respect to the biomass con- sidered as solid fuel. The combustion of a bio-fuel can be better controlled than the combustion of biomass itself, both in terms of operating tempera- tures and of combustion products. With gasification, and with the following clean-up of its products, it is possible to reduce or to remove the content of polluttants and contaminants like N2, Cl and S compounds, for example.
Additionally, storage and transportation of bio-fuels are definitely more ad- vantageous from a logistic point of view [4].
Natural gas (NG) is one of the major primary fuels and its importance is increasing year by year, replacing coal and oil in many applications (indus- trial, residential, but also as fuel for vehicles), and playing a key role in the so-called energy-transition.
In this context, the production of SNG supplies a product of great interest which can be injected in the existing distribution grid and can be used for many different purposes [5].
1.1 Research objective
Purpose of this work is to model a plant based on the integration of a gasifier and of a Solid Oxide Electrolytic Cell (SOEC) for the conversion of woody biomass in synthetic natural gas. The integration of these systems for the conversion of biomass in other chemical products of interest has already been treated by several authors: Pozzo et al. for the prodiction of Dimethyl Ether (DME) [6], Bernical et al. for the production of Fischer-Tropsch products [7], Clausen et al. for the production of methanol [8].
Finally, Giglio et al. studied the integration between biomass gasification and high-temperature electrolysis for synthetic methane production, which is
the matter treated in this work too [9].
With respect to this last work, a different approach was used to model the gasification stage: a kinetic rate model was preferred to the thermodynamic equilibrium one.
A 1D model was realized, which is strictly geometry-dependent and gives a better accuracy in terms of composition of the outlet gas stream from the gasification stage.
The proposed model was validated with respect to the experimental re- sults reported by Campoy et al. [10] and its predictions were compared to the ones reported by Dang et al. [11], showing a good overall accuracy in terms of Round Mean Square Error (RMSE). In this first phase of model vali- dation, a mixture of steam and air was considered as gasifying agent.
Then the reactor geometry was adapted, considering as gasifying agent a mixture of steam and pure oxygen. With respect to the previous valida- tion phase, the input gas superficial velocity was kept constant, leading to a reduction of the geometry of the gasifier.
Subsequently a SOEC was modeled with two main goals: to supply the gasifying agent (oxygen and steam) to the gasifier and, mainly, to enrich the H2content of the syngas produced, to ensure the respect of the stoichiometric requirement for the subsequent methanation reaction. The modeling of this integrated system was performed using APEN Plus software. In particular the gasification stage was modeled with a User defined block “User2”, linked with an external FORTRAN subroutine, used to represent bed hydrodynamics and reaction kinetics. This is the main difference with respect to the previous 0D equilibrium model.
Among all the different types of biofuels which can be produced by the integration of the gasifier and the electrolytic cell, SNG was chosen because methanation reactions are highly exothermic, more than the synthesis reac- tions of other biofuels (for example methanol):
• Methane synthesis reactions
CO + 3H2 → CH4 + H2O ∆h0 = −206 kJ/mol CO2 + 4H2 → CH4 + 2H2O ∆h0 = −165 kJ/mol
• Methanol synthesis reactions
CO + 2H2 → CH3OH ∆h0 = −91 kJ/mol CO + 3H → CH OH + H O ∆h0 = −50 kJ/mol
This is an interesting aspect since the generated heat can be exploited for a thermal integration. About the electrolyzer, the SOEC was chosen for this investigation since it ensures high efficiencies associated with lower costs for hydrogen production [12]. Nevertheless it is not widely used as technology yet, mainly due to some structural problems and materials degradation that can occur due to its high operating temperatures (in the range of 700-900°C usually).
Technology review
2.1 Biomass
It is hard to find a unique definition of biomass. The United Nations Frame- work Conventions on Climate Change (UNFCCC) defined it as:
A non fossilized and biodegradable organic material originating from plants, animals and micro-organisms. This shall also include all products, residues and wastes from agriculture, forestry and related industries as well as the non-fossilized and biodegradable organic fractions of industrial and munici- pal wastes [13].
Unlike fossil fuels, biomass does not take millions of years to grow and develop. For this reason, it is considered one of the most important renewable energy sources.
It can also be considered a "carbon neutral" resource, since its deployment does not add "new"CO2 to the environment. Plants use solar energy to com- bine carbon dioxide and water, converting them into carbohydrates (CH2O)n and oxygen in presence of sunlight, clorophyll and water, in a process called photosynthesis. The sugar formed is stored in a polymer form as cellulose or hemicellulose [14].
nCO + nH O + light → (CH O) + nO
2.1.1 Biomass composition
The biomass main constituents are lignocellulose - which is mainly made up by cellulose, hemicellulose and lignin - and a large number of complex or- ganic compounds. Furthermore, there is an important amount of water stored in form of moisture (M OIST ) and a small amount of inorganic impurities, known as ash(ASH).
The principal constituting elements are carbon(C), hydrogen(H), oxygen (O), nitrogen (N ), chlorine (Cl) and sulphur (S), even if these last three are usually present in very small amounts.
It is very important to know the feedstock composition, in order to deter- mine its potential to produce a valuable product.
The feedstock composition can be given in terms of ultimate and proxi- mate analysis.
• Ultimate analysis
Composition expressed in terms of basic elements, moisture and ash content. It is expressed as weight percentage of the constituent ele- ments, usually on a dry basis. So we can find it expressed as:
C + H + O + N + S + Cl + ASH = 100%
The larger the amount of carbon and oxygen in the feedstock, the greater the amount of CO and CO2 that will be produced during gasification.
At relatively low temperatures, CH4 is produced too, but its content will decrease with increasing temperature, while CO2 and CO content will increase due to the methane reforming and decomposition reactions [15].
• Proximate analysis
It gives the composition in terms of gross components, such as moisture (M OSIT ), Volatile Matter (V M ), ash (ASH) and Fixed Carbon (F C). The volatile matter is the condensable and noncondensable gaseous phase released when the fuel is heated. Its amount depends on some operating parameters, like temperature or the process heating rate.
Ash is the inorganic solid residue left when the fuel is completely burned.
It is mainly constituted by silica, aluminium, calcium and other miner- als. They lead to some problems for the reactor, such as plugging and sintering of the catalysts, especially at high temperatures.
Fixed Carbon (FC) represents the solid carbon in the biomass that re- mains in solid phase after the devolatilization (char).
2.1.2 Biomass conversion
The conversion of solid biomass into secondary products, like liquid and gaseous biomass-derived fuels, can be achieved mainly through two path- ways: biochemical conversion (fermentation) and thermochemical conversion (pyrolysis, gasification).
In this work we will focus on the second pathway, and in particular on the gasification process.
There are several reasons to convert biomass into secondary products [16]:
• secondary products are more energy dense and can be handled more easily: unlike gases or liquids, in fact, solid biomass has great dimen- sions, which makes difficult to store or to transport it;
• the producer gas from a gasifier can be used in a wider range of ap- plications than a solid fuel. It can produce valuable chemicals as side products (methanol, gasoline, fertilizers, etc.);
• flue gases obtained from gasification is less than those obtained from a direct combustion system;
• if the gas obtained from gasification will be used for energy purposes, like combustion in internal combustion engines, SOX and N OX emis- sions are lower, since the syngas is cleaned from its contaminants before being used.
On the other hand, if heat is the only desired product, combustion is preferable and more economical.
2.2 External drying
Woody biomass has a high water content, which could also be up to 50%wt.
on a wet basis. The moisture content depends on several parameters, like the wood type, the harvesting season and the storage locations and conditions, in terms of temperature, humidity of the air, etc.
The external drying of biomass is a very important preliminary stage. Its
to the gasifier or, in general, to the thermochemical process of interest. The biomass water content, in fact, has a great impact on gasification perfor- mances and efficiencies, as it absorbs heat to be evaporated, and that is the main reason why it should be as low as possible [17].
On the other hand, the drying operation has a large impact on the energy balance of the plant overall [18].
The water content is decreased first with a natural drying process, which occurs during the storage period. The speed of this natural process depends on the storage system, in terms of place and duration, the air temperature and the starting water content itself in the stored raw biomass. Usually it is possible to reduce the moisture content up to a value of 30%. If the biomass is stored for too much time – which means over six months – a fraction of organic matter is lost too, due to the natural degradation that occurs. Biomass can lose up to 10-30%of its organic matter per year in this way [17].
To reach values lower than the 30%of moisture left, it is necessary to use industrial dryers, at least to accelerate the process and to avoid the loss of organic matter.
Several technologies are available for the forced biomass drying and they differ one from another depending on some parameters and on the available heat sources. The heat sources of the dryer can be hot flue gases or steam.
The drying medium can operate in a closed loop with a purge to evacuate humidity or can be a once through air stream. Usually, ambient air is heated to lower its humidity and to activate the drying kinetics.
The efficiency of the drying process is less than 50%at low temperatures and it increases with the temperature. At temperatures higher than 160°C, biomass starts to release volatile compounds, and this leads to volatile matter losses and to an increase in the cost of the purification of the drying air [17].
For these reasons it is recommended to keep the drying temperature at values lower than 160°C.
2.3 Gasification
Gasification is the thermochemical conversion of carbonaceous feedstock, like biomass in this case, into a synthetic gas, called syngas, mainly made up by hydrogen(H2)and carbon monoxide(CO), which can be used directly for power generation purposes or for secondary fuels production. Also a residual solid phase, called “char”, is generated, which includes the organic uncon- verted fraction and the inert material (ashes) present in the biomass [19].
There are three main reasons to chose gasification as thermochemical con- version process for biomass:
1. It is possible to increase the produced gas quality, removing non-combustible gases such as nitrogen or water. This increases the heating value of the produced gas;
2. It is possible to reduce the air pollution, removing sulphur and nitro- gen compounds from the produced gas before it is used in combustion engines, for example;
3. It is possible to increase the relative hydrogen content of the fuel.
Gasification needs a so-called “gasifying agent” to occur, which is usually air, oxygen, steam or a mixture of them. In general, it occurs in a oxygen- deficient environment, and this leads to a partial oxidation of the feedstock.
It is possible to recognize three different stages during a gasification pro- cess:
• Drying and devolatilization (or pyrolysis)
It is an endothermic stage. When the biomass is sent to the gasifier, it starts to be heated and releases its moisture content and starts to be decomposed, leading to the formation of lighter molecules. In this stage, three main products are obtained: non-condensable gases - which are mainlyCO,CO2, CH4,H2,H2O and light hydrocarbons -, tars - or liquid fraction, meaning that all these gaseous compounds condense at lower temperatures - and char (Figure 2.1).
This step usually takes place at temperature ranging from 250 to 700°C.
Devolatilization stage can be represented by the following equation [19]:
Biomass → CO+CO2+H2+CH4+H2O(g)+T ars+Char (endothermic)
• Oxidation
It is a very important stage since, being exothermic, it provides the necessary heat to all the other endothermic stages which occur during the gasification. This is the reason why gasification can be considered an auto-thermal process and does not require an external heat source.
Oxidation is carried out in a oxygen-deficient environment or, in other words, with an amount of oxygen which is much lower than the stoichio- metric amount needed for a complete combustion. The main oxidation reactions that occur are:
C + O2 → CO2 ∆h0 = −394 kJ/mol [19]
C + 0.5O2 → CO ∆h0 = −111 kJ/mol [19]
CO + 0.5O2 → CO2 ∆h0 = −283 kJ/mol [19]
H2 + 0.5O2 → H2O ∆h0 = −242 kJ/mol [20]
CH4 + 1.5O2 → CO + 2H2O ∆h0 = −802,31 kJ/mol [20]
• Reduction
It involves all the products of the preceding stages of pyrolysis and ox- idation. All the gas species and the char react with each other, both in heterogeneous and homogeneous reactions, generating the final syngas.
The main reactions which occur are:
C + H2O → CO + H2 ∆h0 = +131 kJ/mol [20]
C + CO2 → 2CO ∆h0 = +172 kJ/mol [20]
CO + H2O → H2 + CO2 ∆h0 = −41 kJ/mol [20]
CH4 + H2O → CO + 3H2 ∆h0 = +206 kJ/mol [20]
The reduction zone temperature has a key role for the determination of the final syngas composition. This because the endothermic reactions are favored at high temperatures, while the exothermic reactions are favored at lower temperatures.
It is important to say that all these stages are not independent and se- quential - even if in this work they were modeled in this way - but actually they occur all at the same time.
2.3.1 Types of gasifiers
Several classifications of gasification reactors can be done on the basis of:
• Gasification agent: mainly air-blown gasifiers, oxygen gasifiers, steam gasifiers;
• Heat source: auto-thermal or direct gasifiers, where the heat is supplied by the partial combustion of the biomass, and allothermal or indirect gasifiers, where heat is supplied by an external heat source;
• Gasifier pressure: atmospheric or pressurised gasifiers;
• Reactor design: fixed bed gasifiers, fluidized bed gasifiers, entrained flow gasifiers.
The following paragraphs about the different types of gasifiers are taken and adapted from Basu [16], which is recommended for further and more detailed informations.
2.3.1.1 Fixed bed gasifiers
These are the simplest kind of gasifiers. The fuel is supported on a grate on the bottom of the reactor, which usually is made moving, for example rotating, and that is the reason why these kind of reactors are also called
“moving bed reactors”. The fuel moves down in the gasifier from the top.
However, due to their configuration, both mixing and heat transfer within the bed are poor, which makes difficult to achieve a uniform distribution of fuel and temperature across the gasifier.
There are three main types of fixed-bed gasifiers:
Figure 2.2: Fixed bed gasifiers: 1. updraft gasifier, 2. downdraft gasifier, 3.
crossdraft gasifier [16].
• Updraft gasifier The gasifying agent enters from the bottom and moves towards the top, while the biomass is loaded from the top and moves downward, so gas and solids are in counter-current mode. The prod- uct gas leaves from near the top of the gasifier. Entering the reactor, the gasifying agent passes through the grate, where it meets the com- bustion zone, with a very high temperature. The hot gas with its low oxygen content moves upward transporting heat to the other zones of the gasifier. As biomass enters from the top, it experiences drying, de- volatilization and gasification, then oxidation. The syngas temperature at the outlet is low, so the tar content is quite high, since the gas does not experience the high temperature required for tar cracking (Fig. 2.2.1).
• Downdraft gasifier It is a co-current reactor, where the gasifying agent enters the gasifier at a certain height below the top. The product gas flows downward, and leaves from lower section of the gasifier through a bed of hot ash. Since it passes through the high-temperature zone of hot ash, the tar in the product gas finds favorable conditions for cracking reactions, so the tar content is very low. An important requirement for this kind of reactor is that the biomass must have a low moisture content, since biomass first meets “low temperature” gases, so it has not heat enough available to be dried, as it happens for example in the updraft gasifier (Fig. 2.2.2).
• Crossdraft gasifier It is a co-current moving bed reactor, in which the fuel is fed from the top while the gasifying agent is injected through a
nozzle from the side of the gasifier. The produced gas is released from the sidewall opposite to the entry point of the gasifying agent, so the gas is produced in the horizontal direction. This means that there is an high temperature in a relative small volume, leading to a very good char conversion and to a very low tar content for the produced gas. Anyway, this solution is not very used (Fig. 2.2.3).
2.3.1.2 Fluidized bed reactors
A fluidized bed is made of fuel particles of specified size and mixed with granular solids of another material called bed particles, which are kept in a semi-suspended condition (fluidization) by the passage of the gasifying agent through them at appropriate velocities from the bottom.
Fluidized-bed gasifiers’ main characteristics are their excellent reactants mixing and temperature uniformity.
A drawback of these reactors is that they usually produce a gas with high particulate content, represented both by the bed material and the solid resid- ual produced from biomass, so a downstream cyclone is necessary as part of the installation.
The fluidized bed design has proved to be particularly advantageous for the gasification of biomass. The gas outlet temperature is usually quite high (about 900°C), so its tar production has been found to be between that for updraft (about 50 g/N m3) and downdraft gasifiers (about 1 g/N m3), with a typical value of 10g/N m3 [21].
• Bubbling fluidized bed gasifier
It was first developed by Fritz Winkler in 1921. It can operate both at low or high temperatures and at atmospheric or elevated pressures. The bed materials are fluidized with steam, air, oxygen or their combination, depending on the choice.
Two phases can be recognized in a bubbling bed gasifier: a dense phase, with an high solid content, located at the bottom of the reactor, and a diluted phase, called freeboard.
In the lower zone gas bubbles formation and implosion generate high turbulence, favouring mixing of the solid components.
The fluidization velocity has to be higher than a certain value called minimum fluidization velocity, which is the velocity which ensures the fluidization of the bed materials. Typical values of fluidization velocities
• Circulating fluidized bed gasifier With respect to BFB gasifiers, it has higher values of superficial gas velocity, usually in the range 3.5-5.5 m/s.
Due to the higher velocities, the solids are dispersed all over the tall riser, allowing a long residence time for the gas and for the fine parti- cles, and this ensures a good carbon conversion.
It is also designed to operate under pressurized conditions in order to increase the final products’ yield (Figure 2.3.2).
Figure 2.3: Fluidized bed gasifiers: 1. Bubbling fluidized bed gasifier, 2.
Circulating fluidized bed gasifier [22].
2.3.1.3 Entrained flow gasifiers
Fine particles and gasifying agent are fed co-currently from the top of the reactor. This results in the gas surrounding or entraining the solid particles as they flow through the gasifier in a dense cloud.
They operate at high temperatures (1300-1500°C) and pressures (25-30 bar) and they are characterized by an extremely turbulent flow which causes rapid and efficient carbon particles conversion (near to 100%).
In general, this solution is preferred more for coal than for biomass, since it is very difficult to grind biomass into such fine dimensions required by this kind of reactor, due to the biomass’ high moisture content.
The high gasification temperatures allow to have a product gas which is nearly tar-free and has a very low methane content.
Furthermore, in entrained flow gasifiers ashes are molten, due to the quite high temperatures which are reached. Molten biomass ash is highly aggres- sive and this reduces the gasifier’s life.
A typical entrained flow gasifier simplified scheme is reported in Figure 2.4.
Figure 2.4: Entrained bed gasifier [23]
2.3.2 Equivalence Ratio (ER) and Steam-to-Biomass Ra- tio (SBR)
Equivalence ratio (ER) and Steam-to-Biomass Ratio (SBR) are two important parameters to define gasifier operating conditions in terms of gasifying agent flow sent to the gasifier.
Equivalence ratio (ER) is defined as the ratio between the mass flow of oxygen supplied as gasifying agent and the theoretical mass flow of oxygen required for the stoichiometric combustion of the biomass.
Equivalence Ratio (ER) = O2 mass f low f ed to the gasif ier
O2 mass f low required f or stoichiometric combustion
As it will be seen later, increasing the mass flow of gasifying oxygen – so ER – leads to an increase of combustion products (CO2 and H2O), which means higher temperature and lower syngas yield and heating value.
Typically ER ranges from 0.2 to about 0.4 [24].
In order to set the gasifying-O2 mass flow to send to the gasifier, it is first needed to know the mass flow of oxygen required for the stoichiometric combustion of that biomass feedstock.
It is known that 1 kg of carbon needs 2.667 kg of oxygen, 1 kg of hydrogen needs 8 kg of oxygen and 1 kg of sulphur requires 1 kg of oxygen to be completely burnt.
Knowing the biomass elemental composition from its ultimate analysis, it is possible to approximate the mass flow of oxygen required for the stoichio- metric combustion as:
O2,stoich = 2.667 · C + 8 · H + S − O
Where C, H, S and O are the mass flows of the corresponding elements in the considered biomass [kg/h].
Steam-to-Biomass Ratio (SBR) is the ratio between the steam mass flow fed to the gasifier and the biomass mass flow fed to the gasifier.
Steam to Biomass Ratio (SBR) = Steam mass f low f ed to the gasif ier Biomass mass f low f ed to the gasif ier
Steam-to-Biomass ratio can be changed by varying biomass mass flow and keeping constant steam mass flow, or vice-versa. Higher SBR values mean both a decrease in the gasification temperature and an increment in the water
content in the produced gas. Also an important energy cost is required to produce more steam, so it is necessary to select an optimal SBR according to different operating conditions.
2.3.3 Tars formation
The term “tar” usually refers to a broad range of organic compounds that are in the form of vapour at the gasification temperature (above 400°C) but liquid at ambient temperature.
Generally, tars are all those organic compounds with molecular weight greater than benzene [25].
They can be classified into primary, secondary and tertiary tars (Figure 2.5).
• Primary tars (or wood oil) are released directly during the devolatiliza- tion step, and mainly depend on the type of biomass. Typical primary tars, as listed by Milne and Evans [21] may be: alcohols, ketones, alde- hydes, phenols, etc.;
• Secondary tars can be formed during the oxidation step, as conse- quence of the increase in temperature that causes the reforming of the primary tars into smaller, lighter non-condensable gases and into a series of heavier molecules. Phenols and olefins are important con- stituents of this group of tars;
• A further increase of temperature leads to the decomposition (crack- ing) and recombination of secondary tars into the so called tertiary (or high-temperature) tars. Tertiary tars start to appear only when primary tars are completely converted into secondary, so they cannot coexist with primary tars. It is possible to further distinguish between:
1. Alkyl tertiary product: including methy derivatives of aromatics, toluene and indene.;
2. Condensed tertiaty aromatics: polynuclear aromatic hydrocarbon (PAH), like benzene, naphtalene, acenaphthylene, etc.
Figure 2.5: Distribution of tar species as a temperature function [21]
In general tar is an highly undesirable product, as it can create many problems to the reactor, including:
• Condensation and subsequent plugging of downstream equipment, where there is a temperature reduction;
• Formation of tar aerosols;
• Polymerization into more complex structures.
Tar in coal gasification comprises benzene, toluene, xylene, etc. and all these compounds have a good commercial value.
Tar from biomass, on the other hand, is mostly oxygenated, due to an higher oxygen content of biomass with respect to coal, and has little commer- cial use.
The tar content has to be reduced, in order to make the produced gas suitable for gas engines applications - and in general for all the other possible downstream utilisations of the syngas - since their presence would lead to a series of problems for the engine and its equipments, as previously said. So tar removal remains an important part of the development and the design of biomass gasification plants.
2.4 Electrolysis
Water electrolysis is the process which allows to convert water and DC elec- tricity into gaseous hydrogen and oxygen, which is the reverse of what an hydrogen fuel cell does.
In a Solid Oxide Electrolytic Cell (SOEC) water acts like a reactant and it is fed from the cathode side of the cell. Oxygen ions are transported to the anode through the electrolyte and the hydrogen in produced cathode side (Figure 2.6).
Figure 2.6: Conceptual design of a SOEC [12]
The overall reaction of water electrolysis, which allows to split water into hydrogen and oxygen supplying electrical and, if required, thermal energy is:
H2O −→ H2 + 12O2
While the semi-reactions which occur at the electrodes are, respectively for cathode and anode side:
H2O + 2e− −→ H2(g) + O2− Cathode O2− −→ 12O2(g) + 2e− Anode
It is easy to notice that the volumetric co-production of oxygen corre- sponds to half of the production of hydrogen.
To avoid re-oxidation of the fuel electrode, a fraction of cathode exhausts can be recirculated, so the inlet stream to the cathode will contain a fraction of hydrogen.
The Reactant Ratio (RR) expresses the reactant molar fraction over the whole inlet molar flow.
Also, not all the reactant will react in the stack. In order to consider this phenomenon, the Reactant Utilization (RU) is used: it indicates which is the percentage of the total flow actually reacting in the stack.
The fraction of inlet molar flow which is actually converted (˙nR):
˙nR = ˙nIN · RU · RR
The overall energy demand of reaction∆H, can be partly supplied by heat (∆Q), while another part (change in Gibbs energy ∆G) has to be supplied electrically:
∆H = ∆G + ∆Q
The minimum electric energy supply required for the electrolysis process is equal to the variation in the Gibbs free energy:
∆G = ∆H − T · ∆S
Where ∆H is the enthalpy variation, T the temperature and ∆S the en- tropy variation.
The overall energy demand ∆H increases slightly with temperature (Fig.
2.7), but the electrical energy demand, ∆G, decreases with the increasing temperature: the ratio ∆G to ∆H is about 93% at 100°C and about 70% at 1000°C [26].
Operating at high temperatures can decrease the cost of the hydrogen produced: since the thermal energy required for the electrolysis reaction can be obtained from the Joule heat produced within the cell and the electrical de- mand is reduced at higher temperatures, the H2 production price decreases.
Figure 2.7: Total, thermal and electrical energy demand for an ideal electrol- ysis process as a function of T [27]
The thermoneutral potential is defined as the potential at which the gen- erated Joule heat in the cell and the heat consumption for the electrolysis reaction are equal:
Utn = z · F∆H
In other terms, it is the voltage at which thermal integration with an exter- nal source is not required and all the energy demand is supplied electrically.
In a real electrolyser the cell voltage for thermoneutral operation is slightly higher thanUtn due to heat losses and thermodynamic irreversibilities. Typi- cally, for a SOEC, this voltage is of around 1.29 V.
Operating below this value, the heat must be supplied to the cell to main- tain the temperature (endothermic mode), while operating above this volt- age, the cell operates in exothermic mode and the produced heat has to be extracted.
If H2O is fed into the system in liquid phase, also the heat demand for water evaporation has to be considered and there will be an increase in the operation voltage [26]:
Vvap = ∆Hz · Fvap
The thermoneutral potential is then defined as the sum of both contribu- tions:
Utn,liq = ∆H + ∆Hz · F vap
The possible high heat utilisation of internal losses is one of the major motivation to operate high temperature electrolysis (at 700-900°C).
In general, low temperature electrolysers are operated above the ther- moneutral voltage due to high internal losses or overvoltages. This results in a heating of the electrolysis cells requiring external cooling of the module.
For high temperature electrolysers, instead, thermoneutral voltage rep- resents the standard operation mode. The cell is operated at constant tem- perature as internal heat production by irreversibilities is equalised by heat consumption of the electrolyser reaction.
2.4.1 Polarization curve
The relationship between voltage and current density is a fundamental char- acteristic of the cell efficiency.
The cell voltage can be expressed as the sum of the reversible cell voltage Urev, also called Open Circuit Voltage (OCV) and the overvoltages caused by ohmic resistance Uohm, limitations in electrode kinetics (activation overvolt- agesUact) and mass transport (concentration or diffusion overvoltageUdif f):
U = Urev + Uohm + Uact + Udif f
First of all, the OCV (or Urev) is the difference of potential between the two terminals of the cell when it is disconnected from any circuit. It can be estimated with the Nernst equation:
OCV = z · F∆G
Where ∆G is the Gibbs free energy variation, z the number of electrons involved in the reaction (which is 2, in this case) and F the Faraday constant.
As previously written, the operating voltage differs from OCV due to some irreversibilities which cause overvoltages within the cell:
• Activation overvoltage (ηact). It is due to two phenomena: the chem- ical equilibrium state of ions at the electrode-electrolyte interface and the overcoming of the electric field due to transfer of charged particles across the interface by ions. It can be decreased increasing tempera- ture, active surface area or activity of the catalyst;
• Ohmic overvoltage (ηohm). It is caused by the resistance to conduction of ions in the electrolyte, conduction of electrons in the electrode and contact resistance. Normally, only ionic resistance are considered, all the others are neglected;
• Diffusion or concentration overvoltage (ηdif f). It is very relevant at high current densities, due to transport phenomena;
All the overvoltages previously described grow with the current density, but one of the mechanisms prevails over others depending on the operating zone.
The dependency between cell voltage and current (or current density) is shown in Figure 2.8 with a review of different SOEC characteristics from liter- ature [27]. The current-voltage (I-U) curve characterises the electrochemical behaviour of an electrolysis cell.
Figure 2.8: Overview of characteristic I-U curves of SOECs from literature
A simplest model can be obtained assuming linear relationship between voltage and current density (linearization of the polarization curve).
This hypothesis is well justified for a SOEC since high temperatures en- hance the kinetic of electrochemical reactions, increasing the reaction rate and decreasing the activation overvoltage, and improve the diffusion of reac- tants in electrodes, decreasing the diffusion overvoltage.
The slope of this first order curve is the Area Specific Resistance (ASR), defined as:
ASR = VC− OCVj
WhereVC is the cell potential [V], j is the current density [A/m2].
So the operating voltage can be expressed, as a function of current den- sity in this way:
VC = OCV + j · ASR
ASR is influenced by materials used for the construction of the cell (both electrolyte and electrodes), geometrical features and operational parameters, like temperature, pressure and inlet gas composition [28].
2.4.2 Cell efficiency
The electrical power needed by the cell can be calculated as:
Pel = VC· j · S
Where VC is the cell potential, j the current density and S the active sur- face area of the cell [m2].
The current density is proportional to the hydrogen production rate, as expressed by the Faraday’s law, considering also the Faraday efficiency ηF (or current efficiency) which is defined as the ratio between actual hydrogen production rate and the theoretical one. For the modules, withncelectrolysis cells and operating at current I, the total hydrogen production rate in Nm3/h is then:
V˙H2 = ηF · (n2·FC·I) · [2.414 · 3.6N mmol3 · hs]
And the efficiency of an electrolyser can be defined as
ηHHV = V˙H2·HHVP H2
el
Where HHVH2 is the Higher Heating Value of hydrogen (3.54 kWh/Nm3) andPel the electric energy consumption in kW. The cell voltageUC given in V is inversely proportional to the efficiency:
ηHHV = ηF·nC ·I2·Fn ·HHVH2
C·UC·I
The electrolyser efficiency could also be referred to the lower heating valueLHVH2(3.00 kWh/Nm3):
ηLHV = V˙H2·LHVP H2
el = 3.003.54 · ηHHV = 1.25 VU
C·ηF
It is also possible to define the specific energy consumption for the produc- tion of 1 Nm3 (or 1 kg) of hydrogen, which is proportional to the cell voltage:
ES = LHVη H2
LHV = η2.4VF·UC
The efficiency of an electrolyser decreases:
- with increasing current density (and increasing Ucthen);
- with decreasing temperature;
- slightly with increasing pressure.
2.4.3 Solid Oxide Electrolytic Cell (SOEC)
It is an electrolytic cell which operates at temperatures of 700-900°C. High temperature means higher efficiencies, but also that great attention has to be paid to the material stability. A simplified process layout of a Solid Oxide Electrolysis (SOE) system is shown in Figure 2.9 [27].
The feed water or steam is pre-heated in a recuperator against the hot product streams leaving the stack. Low temperature heat has to be inte- grated to account for the heat of evaporation. The stack consists typically of planar cells electrically connected in series. Steam and recycled hydrogen are supplied to the cathode to maintain reducing conditions and are partly converted to hydrogen. The mixture of steam and hydrogen is eventually separated by cooling and condensing the water.
Figure 2.9: Simplified layout of a SOE system [27]
Air can be used at the anode as “sweep stream” for the removal of product oxygen. As demonstrated by Barelli et al. [29], it is possible to substitute air with steam on the oxygen electrode side as sweep gas, without any decrease in performance. It also gives the opportunity to produce pure oxygen without any further separation cost needed, since it is only necessary to condense water to separe it from oxygen, while if air is used as “sweep unit” as Air Separation Unit is necessary.
The sweep gas is needed because the anode, which is the oxygen elec- trode, presents the higher contributions to ohmic losses during the operations and the higher degradation due to delamination of the electrode/electrolyte interface. Also, high oxygen concentrations in the electrode increases polar- ization losses. So the function of the sweep gas is to “clean” the electrode flowing out the oxygen from the anode, reducing in this way the oxygen con- centration.
In this work steam was preferred as sweep gas, since it is already available in the system for the production of hydrogen in the SOEC and because the gasifying agent needed in the gasification stage is a mixture of oxygen and steam.
2.4.4 SOEC materials
The most common electrolyte material is a dense ionic conductor, consisting of ZrO2 doped with 8 mol% of Y2O3 (YSZ). This material presents high ionic
conductivity as thermal and chemical stability at the operation temperatures, which are of about 800-1000°C. For the fuel electrode (cathode), the most commonly used material is a porous cermet composed of YSZ and metallic nickel. While, for the oxygen electrode, the most used material is the lan- thanum strontium manganite (LSM)/YSZ composite [26].
2.5 Syngas clean-up system
The raw syngas produced in the gasifier is rich of contaminants, mainly tars and other inorganic compounds, like nitrogeneous compounds (e.g. NH3), sulfur containing compounds (e.g. H2S) and hydrogen halogens (e.g. HCl).
Before being used in downstream applications, the raw syngas has to be cleaned then with different levels of purity required depending on the final purpose. There are two main possibilities to reach this goal: low temperature processes (cold gas cleanup) and mid-to-high temperature processes (hot gas cleanup) [30].
2.5.1 Cold gas clean-up
Cold gas clean-up is considered the conventional approach for syngas clean- ing, due to its high contaminant removal efficiency. It is carried out at low temperature, which means at or even below ambient temperature.
The main disadvantage of this approach is the loss of efficiency due to the syngas cooling, since gasification is usually carried at high temperatures, in the range of 800-1200°C.
Cold gas cleanup utilizes “wet” or “dry” processes.
• Dry cold gas clean-up processes employ mechanical, physical and elec- trostatic separators, like cyclones, electrostatic precipitators and other filters.
• Wet cold gas clean-up processes, on the other hand, employ mainly spray and wash towers: these units remove contaminants by absorption and adsorption, filtration or a combination of them.
Wet cold gas cleanup processes are most commonly used, since they allow for multiple-contaminants removal. It is a very suitable solution especially for the inorganic contaminants. For example HCl, NH3 and H2S are very soluble in water, so wet towers or scrubbers using water as a solvent will remove all
in water. Further and more detailed informations about the cold gas clean-up methods and processes have been reported by Abdoulmoumine et al. [30].
2.5.2 Hot gas clean-up
Hot gas clean-up approach is of great interest in contaminant removal from syngas. It is carried out at temperatures higher than 300°C.
The main advantage of this solution is that it does not require to cool down the raw syngas, thus avoiding a loss in efficiency. It also reduces the waste streams with the potential of converting some contaminants into useful products [30].
There are two different solutions to reduce the tar content in syngas: pri- mary methods, which are the methods operating directly during the gasi- fication stage to limit tars formation, and secondary methods, also known as post-gasification methods, which operate downstream of gasification. Sec- ondary methods involve physical, thermal or catalytic treatments [31].
Tar compounds can be removed mainly by thermal cracking and steam or dry reforming. Thermal cracking is the conversion of tar species into C and H2 and it requires temperatures higher than 1100°C.
CnHm −→ nC + m2H2 T hermal cracking
Tar as well as light hydrocarbons can be converted to carbon monoxide (CO) and hydrogen (H2) through steam reforming, as shown below
CnHm+ nH2O (g) −→ nCO + n+m2 H2 Steam ref orming
Steam reforming of tars occurs only above 900°C in the absence of a cat- alyst [30].
In any case, both these two solutions are better performed with the em- ployment of a catalyst, directly into the gasifier.
The right choice of the catalyst is an important task. A good primary catalyst should have good activity for the conversion of tars into valuable gaseous products (e.g. H2 and CO), but also good stability and resistance to attrition and deactivation through coking, sulfur poisoning or sintering [31].
Generally tar cracking and steam reforming reactions, as well as partic- ulate removal, which is not considered in this model, are obtained in down- stream components, like ceramic catalytic filters (or candles). These com- ponents might also be positioned in the freeboard of a BFB reactor, having the possibility to obtain a very compact unit. In this way it is also possible to combine hot gas filtration and cleaning from inorganic impurities with cat- alytic tar reforming in a single stage. Some of the used catalysts, in fact, act also as a primary sorbent for ammonia, hydrogen sulphide and hydrogen chloride. The mineral catalyst like dolomite (CaO-MgO) and olivine (Mg2SiO4- Fe2SiO4) and the nickel and iron based catalysts have been proven to be the most effective at temperatures raging from 500 to 900°C. [30].
Great attention has to be paid to other inorganic contaminants removal, since they are responsible for corrosion and other issues especially for the downstream equipments. It is important to consider one or more blocks to remove or reduce the contaminants concentration below the limits required for the final applications. The quasi-total removal of hydrogen sulphide at temperatures of about 400°C is obtained thanks to solid-gas reactions be- tween the contaminant and a metal oxide (MO), especially Zn, Cu and Ce oxides. HCl is treated with alkali-based sorbents, like NaHCO3, in the tem- perature range of 526-650°C. In both cases it is possible to reach contaminant concentrations of 1 ppm or lower [32].
2.6 Methanation
Methanation is a physical-chemical process which aims to produce methane from hydrogen and carbon oxides.
The two methanation reactions are:
CO + 3H2 −→ CH4 + H2O (g) ∆h0 = −206 kJ/mol CO methanation CO2 + 4H2 −→ CH4 + 2H2O (g) ∆h0 = −164 kJ/mol CO2 methanation
Both reactions are exothermic. For each 1 m3 of methane produced per hour, 2.3 and 1.8 kW heat are produced respectively [33]. CO2 methana- tion is a linear combination of CO methanation and reverse water-gas shift (RWGS) reaction, which always accompanies the CO methanation reaction using nickel catalysts.
CO2 + H2 −→ CO + H2O (g) ∆h0 = 41kJ/mol
The equilibrium of both reactions is influenced by pressure and tempera- ture [33].
In thermodynamic equilibrium, high pressures favor the production of methane. High temperatures, on the other hand, limit methane formation.
However, the performance becomes poor at temperatures below 200°C and too high pressures, because of reaction kinetics limitations.
The input gas for the methanation section must have a composition with the correct ratios between the reactants, which are CO, CO2 and H2.
The predominant methanation reaction is the CO methanation, which re- quires an H2/CO = 3. But also CO2 methanation has to be taken into account, so the feed gas has to meet the following stoichiometric requirement, which is called “Feed Gas Module”:
F EED = yyH2− yCO2
CO+ yCO2 = 3
Where yi represents the molar fraction of the reactant i.
Catalytic methanation reactors are typically operated at temperatures be- tween 200°C and 550°C and at pressure raging from 1 to 100 bar [34]. A catalyst is required to achieve acceptable rates and high selectivity. Several materials may be used, but the most common choice is represented by Ni, due to its availability and low cost. Nickel based catalysts require a high purity of the feed gas, so great attention has to be paid to syngas cleaning before it reaches the methanation section: in particular it is necessary to remove sulphurous compounds, which are responsible of sulphur poisoning. In this work, sulphur compounds are represented only by H2S.
Methanation is a strongly exothermic reaction, so another big issue is to re- alise a good temperature control in the reactor, since high temperatures can cause catalyst sintering and carbon deposition. The chosen design consists in a series of cooled reactors with an intermediate condensation stage in order to condense and remove the produced water.
Process modelling
3.1 General aspects
The commercial software ASPEN Plus was used to model the plant. It is a process modelling tool for conceptual design, optimization and performance monitoring of chemical processes.
It has a huge database of pure components and phase equilibrium data for conventional chemicals, electrolytes, solids and polymers and with reliable thermodynamic data, so it is possible to simulate the actual plant behavior, using some engineering relationships, physical and chemical laws, such ass mass and energy balances, phase and chemical equilibrium and reaction ki- netics.
In this work, the hypothesis of chemical equilibrium was applied to all the reactors, but the User2 block which represents the gasification stage [35, 36, 37].
The stream class MIXCINC was selected. Stream classes are used to define the structure of simulation streams. The selected one includes:
• Conventional streams (MIXED), i.e. vapour/gas and liquid streams and solid salts in solution chemistry;
• Conventional inert solids (CI), i.e. solids that are inert to phase equi- librium and salt precipitation/solubility, but take part to chemical equi- librium (using Gibbs reactors);
• Non-conventional solids (NC), i.e. heterogeneous substances inert to phase, salt and chemical equilibrium that cannot be represented with a
To calculate non-conventional solids properties, it is necessary to specify the “NC Props”. The only physical properties needed in this case are enthalpy and density.
This is done using the HCOALGEN and the DCOALIGT models. In partic- ular, the HCOALGEN uses the proximate analysis, the ultimate analysis and the sulphur analysis to calculate the enthalpy of the non-conventional solid considered.
All the values of proximate, ultimate and sulphur analysis are defined as weight percentage on a dry basis, with the exception of moisture in proximate analysis, which is expressed clearly on wet basis.
The sum of the values of the sulphur analysis must be equal to the value for sulphur specified in the ultimate analysis.
The value specified for ash must be equal for ultimate and proximate anal- ysis.
The sum of the values specified for ultimate and for the proximate analysis must be equal to 100 in both cases.
All the components used in this work were specified:
• BIOMASS and ASH as NC;
• C as conventional inert solid;
• H2O, CO, CO2, CH4, H2, O2, C3H6O2, C7H8, C6H6, C6H6O, C10H8, NH3, H2S, HCl, N2 as conventional streams.
Finally, the thermodynamic methods used to calculate the properties of the conventional streams were selected:
• The IDEAL property method (Ideal gases, Raoul’s Law and Henry’s Law) was selected for the drying and gasification section;
• The PENG-ROB property method (Peng Robinson equation of state) was selected for all the other sections.
A simplified process flowchart is reported in Figure 3.1.
Figure 3.1: Simplified process flowchart
